Synthesis of organic compounds



Feb. 20, 1951 G MCGRATH 2,542,422

I SYNTHESIS OF ORGANIC COMPOUNDS Filed Oct. 31, 1947 2 Sheets-Sheet l 1N VENTOR. HENRY G M0 GRA TH E I W ATTORNEYS Patented Feb. 20, 195i SYNTHESIS OF ORGANIC COMPOUNDS Henry G. McGrath, Elizabeth, N. J., assignor to The M. W. Kellogg Company, Jersey City, N. 1., a corporation of Delaware Application October 31, 1947, Serial No. 783,382 In Great Britain July 25, 1947 19 Claims.

This invention relates to the synthesis of organic compounds. In one aspect this invention relates to the hydrogenation of carbon monoxide in the presence of ametal hydrogenation catalyst to produce hydrocarbons and oxygenated organic compounds. In another aspect this invention relates to the starting-up procedure for the hydrogenation of carbon monoxide in the presence of a catalyst comprising iron. The present process is applicable also in reacting hydrogen with other organic compounds containing the carbonyl group and herein designated as carbon oxides," such as carbon dioxide, ketones aldehydes, acyl halides, organic acids and their salts and esters, acid anhydrides, and amides. In the following description of the invention, the hydrogenation of carbon monoxide will be referred to specifically. It will be understood, however, that theinvention is of wider application including in its scope the hydrogenation of any suitable carbon oxideor mixtures thereof.

This application is a continuation-in-part of my prior and copendingapplication Serial No.

. 725,835, filed February 1, 1947, and application Serial No. 69 820, filed August 15, 1946. In application Serial No. 690,820 I am co-inventer with Louis C. Rubin. These applications, in general, relate to the synthesis of organic compounds in the presence of a fluidized suspended hydrogenation catalyst. In application Serial No. 690,820 thesta'rting up procedure of the present invention isdisclosed in" the examples.

f-Iroh, as well as other metals of group VIII of the periodic table, has been used as a catalyst for the hydrogenation of carbon monoxide. A suitable metallic iron catalyst is prepared by various methods, such as by precipitation and the subsequent reduction ,of a ferruginous precipitate to elementary iron or by. the direct reduction of naturally occurring magnetite to elementary iron. A cheap iron catalyst is that obtained by fusing Alan Wood Ore and the subsequent reduction of the fused material. The method of preparation of an iron catalyst from Alan Wood Ore is described in detail in my prior and co-pending application Serial No. 735,536, filed March 18, 1947.

(Cl. Milk-449.6)

It has been found, however, that although these catalysts are capable of high activity and the production of relatively large quantities of normally liquid organic compounds by the hydrogenation of carbon monoxide, in almost all instances the freshly reduced iron catalyst is substantially inactive for producing normally liquid products and is further characterized by being diflicult to fluidize at operating conditions when used in. finely divided form. The freshly reduced metal catalyst produces only a relatively small quantity of normally liquid hydrocarbons, that quantity often being substantially less than about 25 or .30 cos. per cubic meter of fresh feed, and the fresh catalyst exhibits a carbon monoxide conversion of less than about 45 per cent. Furthermore, after a short period of operation the finely divided iron catalyst often tends to agglomerate and, when the catalyst is fluidized by suspending in a gas to form a so-called pseudo-liquid dense phase, the dense phase settle or becomes deaerated. A settled or deaerated condition, regardless of the reason for settling, is characterized by channeling and rat holing of the reactant gases through the catalyst mass. According to thisinvention, it has been found that such freshly reduced catalyst must be activated prior to establishing operating conditions in a particular manner hereinafter described in order to achieve maximum activity and stability of the catalyst for the production of normally liquid organic compounds without excessive formation of wax and relatively high-boiling organic compounds, and in order to prevent agglomeration and sticking of the fluidized catalyst.

It is an object of this invention to provide a process for the production of organic compounds 7 by the hydrogenation of carbon oxides.

Another object oithis invention is to provide a process for the hydrogenation of carbon monoxide to produce hydrocarbons having more than one carbon atom per molecule and oxygenated organic compounds.

It is still another object of this invention to increase the yield of normally liquid hydrocarbons by the hydrogenation of carbon monoxide in the divided form for the hydrogenation of a carbon oxide to produce organic compounds.

Another object of this invention is to provide a method for starting-up or initiating the reaction between hydrogen and carbon monoxide in the presence of a freshly reduced or prepared iron catalyst.

A still further object of this invention is to provide a method for maintaining an iron catalyst at its maximum activity during the process for the hydrogenation of carbon monoxide by the fluidized technique.

It is still another object of this invention to provide an integrated process for the hydrogenation of carbon monoxide by the fluidized technique and the simultaneous and continuous or intermittent regeneration of a portion of the fluidized catalyst.

Another object of this invention is to provide a method for amuring continuous operation of the hydrogenation of carbon monoxide in the presence of a fluidized iron catalyst.

Yet another object is to provide a method for rendering an iron catalyst susceptible to severe operating conditions without danger of loss of fluidity of the catalyst.

Another object is to provide a method for prevention of settling of the catalyst bed in the hydrogenation of carbon monoxide by the fluid-bed technique.

Various other objects and advantages of the present invention will become apparent to those skilled in the art from the accompanying description and disclosure.

In the hydrogenation of av carbon oxide in the presence of a catalyst according to this invention, the freshly reduced or freshly prepared catalyst is subjected to a preliminary treatment prior to establishing the operating conditions necessary to produce the maximum yield of desired organic compounds, particularly normally liquid organic compounds. It has been found that a freshly prepared or freshly reduced catalyst. comprising iron, such as that prepared by the fusion of a naturally occurring magnetite and the subsequent reduction of the fused material to produce elementary iron, is relatively inactive as a catalyst for the hydrogenation of carbon monoxide and, when used in a fluidized condition, the catalyst has a tendency to lose its fluidity, such as by agglomerating, sticking, channeling, etc. In order to overcome these limitations encountered with a freshly reduced or prepared catalyst, the catalyst must be brought to thedesired operating conditions by a pretreatment or an activation treatment in accordance with the teachings of this invention. Pretreatment of the freshly reduced or prepared iron catalyst constitutes, according to this invention, initiating contact between synthesis gas comprising hydrogen and carbon monoxide in a mol ratio between about 1:1 and about 4:1 and having a carbon monoxide concentration of at least volume per cent at a pressure below about 90, preferably below about 85, pounds per square inch gage and thereafter increasing the pressure above about 100 pounds per square inch gage to the desired operating pressure for effecting the synthesis of organic compounds.

The pressure is increased from below about pounds per square inch gage to pounds per square inch gage during contact of at least about 200 standard cubic feet of synthesis gas of the 5 above composition with each pound of iron (Fe) at a space velocity not greater than about 40 standard cubic feet of gas per pound of iron (Fe) per hour. Actually, the limitation as to space velocity is the capacity of the system to remove the exothermic heat of reaction in order to prevent an excessive rise in temperature. Preferably, the pressure is raised to about pounds per square inch gage during contact of at least 300 standard cubic feet of synthesis gas with each pound or iron (Fe) in the reactor.

In the event the desired operating pressure is above 100 or 150 pounds per square inch gage, the pressure is raised gradually from the latter pressure of 100 or 150 pounds per square inch to the operating pressure, preferably at a rate not in excess of about 50 pounds per square inch gage per 100 standard cubic feet of synthesis gas contacted with each pound of catalyst. In the event the desired op.rating pressure used for effecting the-hydrogenation of carbon monoxide to produce the desired product is below about 100 pounds per square inch gage, the pretreatment is carried out until the pressure is above about 100, preferably above 150,. pounds per square inch gage in the manner previously described, and then the pressure is subsequently reduced to the desired operating pressure.

The initial temperature of contact between synthesis gas and fresh catalyst is at least 450 F., preferably above 500 F. for a low alkali catalyst and above 550 F. for a high alkali catalyst to prevent the condensation of wax and relativcly high-boiling organic compounds on the iron catalyst particles. The temperature during the pretreatmtnt is increased from the initial temperature of contact to the desired operating temperature or above and for best results is raised above about 600 F. Raising the temperature during the pretreatment may be conveniently achieved simultaneously with the operation of raising the pressure, although the temperature first may be raised to a predetermined level and then the pressure raisid, or vice versa. Generally, the temperature of reaction is maintained below about 700 F. because above such temperature the formation of coke or carbon on the catalyst is rapid. Coke formation on the catalyst decreases the catalyst density which in turn causes an expansion of the fluid catalyst bed to such anextent that the catalyst may often pass from the reactor through the gaseous efiluent outlets in the upper portion of the fluid-bed'reactor. After the pretreatment has been carried out and the operating temperatures and pressur reached, the yield of normally liquid organic compounds is substantially at its maximum and is usually in excess of about 80 005. per cubic meter of fresh feed gas.

As an example of the present invention, the desired operating conditions may be a pressure of about 250 pounds per square inch gage and a temperature of about 580 R, such operating conditions being chosen upon consideration of such factors as product, economics and mechanical limitation of the equipment. A freshly reduced fused iron catalyst having a relatively low alkali content below about 0.7 weight per cent calculated as the oxide is initially contacted in a fluidized condition initially with synthesis feed gas 75 comprising hydrogen and carbon monoxide in a moi ratio of about 2:1 at a temperature of about 450 F. and at a pressure of about 15 or 20 pounds per square inch gage. The pressure and temperature are raised to about 150 pounds per square inch gage and about 650 F., respectively, during the time required for contact of about 350 standard cubic feet of synthesis gas with each pound of catalyst (Fe) in the reaction zone. When the pressure has reached about 150 pounds per square inch gage the catalyst exhibits high activity characterized by a contraction of gases of about 50 per cent in the reaction zone and yields of oil and water of about 60 cos. and 70 cos. per cubic meter of fresh feed, respectively. The pressure is then raised to a predetermined operating pressure of about 250 pounds per square inch gage and the temperature reduced to a predetermined operating temperature of about 580 F. for producing the maximum yield of normally liquid organic compounds withoutexcessive formation of wax or relatively high-boiling organic compounds.

When the catalyst contains a relatively high alkali content (above about 0.8 per cent of an alkali metal or an alkaline earth calculated as the oxide), the amount of synthesis gas contacted with the fresh catalyst below 100 pounds per square inch pressure is preferably greater than about 250 standard cubic feet per each pound of catalyst in the reaction zone. For a more detailed description of the composition of high and low alkali catalysts, reference may be had to my prior and co-pending application Serial No. 725,835, filed February 1, 1947.

As used herein, operating conditions refer to those conditions used after pretreatment or acti-vation of the catalyst for producing the desired product of the process and are usually substantially constant in contrast to the changing conditions during pretreatment of the catalyst.

According to the preferred embodiment of thislnvention during the pretreatment or activation period as well as after the iron catalyst has been pretreated and the main operating period started, a aseous mixture comprising hydrogen and a carbon oxide is passed through a reaction zone in contact with a suspended mass of the finely divided iron catalyst. The gaseous mixture of reactants and reaction products is passed through the mass of finely divided catalyst at a linear gas velocity sufiicient to suspend or entrain the catalyst mass in the gas stream. Generally, the gas stream is passed upward through the reaction zone at a linear velocity sufiiciently low to maintain the catalyst in a dense fluidized pseudo-liquid condition. However, the velocity maybe sufficiently high to entrain all or a substantial proportion of the finely divided catalyst in the gas stream to form a continuous catalyst phase which circulates with the flowing gas stream without departing from the scope of this invention. In the former condition in which the gaseous mixture is passed upward through the catalyst mass, the catalyst mass may be said to be suspended in the gas stream but not entrained therein in the sense that there is movement of the catalyst mass as such in the direction of flow of the gas stream. When operating with a catalyst in the pseudo-liquid condition, it is preferred to maintain the upward velocity of the gas stream sufficiently high to maintain the fluidized catalyst mass'in a highly turbulent condition in which the catalyst particles circulate at a high rate within the pseudo-liquid mass. In this pseudo-liquid condition of operation a 6 small proportion of the catalyst in the fluidized mass may be entrained in the gas stream emerging from the upper surface of the gas stream whereby catalyst thus entrained is carried away from the mass. In producing normally liquid organic compounds under conditions to react all or a major proportion of the carbon monoxide reactant by employing a finely divided iron catalyst suspended in a pseudo-liquid condition in the reaction zone, the preferred specific operating conditions are approximately as follows: a pressure between about atmospheric and about 600 pounds per square inch gage, a temperature between about 350 F. and 700 F. and a space velocity equivalent to a charging rate between about 100 and 5000 volumes of combined or total feed gas per hour per volume of catalyst in the dense phase, more preferably a' space velocity above about 500 or 600.

The catalyst material towhich the present invention applies is a finely divided powder comprising reduced metallic iron'and may contain in addition appropriate amounts of a promoter or promoters incorporated with the iron in the manner described in the aforementioned copending application Serial No. 735,536, or a mixture of such iron catalyst and other catalytic materials and non-catalytic materials. The catalyst may also include in combination therewith supporting materials, such as alumina, silica gel, bentonite type clay, etc. In this specification and claims, the catalyst is described by reference to its chemical condition subsequent to its reduction and prior to pretreatment.

The catalyst is employed in a, fine state of subdivision. Preferably, the powdered catalyst initially contains no more than a minor proportion by weight of material whose particle size is greater than 250 microns. Preferably, also, the greater proportion of the catalyst mass comprises material whose particle size is smaller than microns, including at least 25 weight per cent of the material in particle sizes smaller than 40 microns. A highly fiuidizable powdered catalyst comprises at least 75 per cent by weight of material small than 150 microns in particle size, and at least 25 per cent by weight smaller than about 40 microns in particle size Operations in which the above finely divided catalyst is maintained in a pseudo-liquid fluidized condition in the reaction zone results in a concentration or density of catalyst expressed as pounds per cubic foot between one quarter and three quarters of the density of the catalyst in a freely settled condition. For example, with finely divided reduced iron, the freely settled density is about to about pounds per cubic foot and the density of the pseudo-liquid dense phase is between about 30 and about 105 pounds per cubic foot depending upon the condition of the catalyst as to coke, wax, etc. In contrast the concentration or density of an entrained finely divided catalyst in a high velocity system, is less than about one sixth of the freely settled density of the catalyst, and for reduced iron is often about 10 or 12 pounds per cubicfoot.

With the catalyst present in a pseudo-liquid condition, the powdered catalyst mass is maintained in a reactor substantially larger than the volume occupied by the catalyst mass in the fluidized condition. In this operation all but a minor proportion of the catalyst mass is contained in the dense fluidized pseudo-liquid mass, which may be designated as the dense phase of the catalyst. The dense phaseof the catalyst occupies 7 the lower part of the reactor while that part of the reactor above the dense phase is occupied by a mixture of gases and powdered catalyst in which the catalyst concentration is much lower, and of a different order of magnitude, than the concentration of the catalyst in the dense phase. This diffuse phase may be said to be a disengaging zone in which the solids lifted above the dense phase by the gas stream are disengaged there- 'from and returned to the dense phase to the extent that such solids are present in the dif-' mum in the upper part of the reactor. Between,-

the dense phase of high average concentration and the diffuse phase of low average concentration there is a relatively narrow zone in which the concentration of solids in the gas stream changes in a short space from the high concentration of the dense phase to the low concentration of the diffuse phase. This zone has the appearance of an interface between two visually distinct phases. The dilute or diffuse phase generally has a concentration of catalyst less than about 0.01 pound of catalyst per cubic foot of gas.

This dense phase operation ordinarily involves employment of catalyst powders and linear gas velocities such that a relatively small portion of the dense fluidized catalyst moss is carried away by entrainment, and it is necessary, therefore, to provide means in the reactor for separating such entrained catalyst and returning it to the dense phase, or to provide means externally of the reactor to separate entrained catalyst from the gaseous effluent and return it to the reactor, or otherwise to recover catalyst from the gaseous efiluent.

When catalyst is permitted to pass out of the reactor by entrainment in the gas stream in either the pseudo-liquid operation or the continuous phase operation, it is necessary to return such catalyst to the reactor, or replace it with'fresh or revivifled catalyst, in order to maintain the desired volume of fluidized catalyst in the reaction zone.

The linear velocity of the gas stream passing upward through the dense phase is conveniently expressed in terms of the superficial velocity, which is the linear velocity the charge gas stream would assume if passed through the reactor in the absence of catalyst. This superficial velocity takes into account the shrinkage in volume caused bythe reaction and is, preferably, in the range of from about 0.1 to about 6 feet per second. When operating with a continuous catalyst phase in which the catalyst is entrained in the flowing gaseous mixture, velocities as high as about 50 feet per second may be used. Reference may be had to my prior and co-pending application Serial No. 726,620, filed February 5, 1947, for a more detailed discussion of velocities and other conditions characteristic of a-high velocity system.

In hydrogenating to the preferred embodiment of this invention, unconverted hydrogen and/or carbon monoxide carbon monoxid according are recycled in a ratio of combined or total feed to fresh feed above about 1:1 to about 5:1. The recycle gas containing hydrogen and/or carbon monoxide may additionally contain normally gaseous and/or normally liquid components of the reaction efiluent. An overall conversion of carbon monoxide between about 70 and about 100 per cent has been observed when using such recycle ratios. A yield of oilbetween about 97 and about 126 ccs. per cubic meter of fresh feed gas and a yield of water between about 99 and about 171 ccs. per cubic meter of fresh feed gas have been obtained under the preferred operating conditions of this invention with an activated or pretreated catalyst.

The feed mixture to the reaction zone, as previously stated, comprises hydrogen and-carbon monoxide in the previously indicated ratios. In most instances relatively non-reactive ingredi-- ents may also be contained in the feed mixture, such ingredients comprise nitrogen, methane, ethane and other saturated hydrocarbons. Since the feed gas contains other such ingredients as nitrogen, it is necessary inmost instances to vent or discard a portion of the recycle gas in order to prevent a build-up of nitrogen in the system.

Upon extended and prolonged use of the iron catalyst in the hydrogenation of carbon monoxide, it becomes necessary to regenerate or revivify the catalyst as a result of accumulation of carbonaceous deposits thereon and as a result of deactivation of the catalyst itself. Carbonaceous deposits may be removed from the catalyst and the catalyst reactivated by treating the catalyst by the successive steps of oxidation and reduction, or reduction alone, accompanied by a subsequent activation treatment similar to that used for the starting-up procedure or pretreatment with the freshly prepared or reduced iron catalyst. Oxidation of the finely divided catalyst is conveniently carried out by the fluidized technique with an oxygen-containing gas, such as air, at the operating pressure of the synthesis reaction, or at a lower pressure, and at a temperature above about 800 F. Generally, the pressure for oxidation is approximately atmospheric. Reduction is normally carried out at substantially lower temperatures than the oxidation temperature when super-atmospheric pressures such as above 200 pounds per square inch gage are used; such reduction temperatures may be as low as about 600 to about 800 F.

When reduction of the catalytic material is effected at atmospheric pressures, the temperature is between about 900 F. and about 1600 F. The reducing gas preferably comprises gaseous hydrogen; however, other reducing gases, such as methane, carbon monoxide, etc., may be employed and the gas stream may include other non-reducing ingredients such as nitrogen, in

amounts which do not interfere with the reducing action.

The fiuidization of the catalytic material for both the oxidation and reduction operations may be brought about initially by the passage of the stream of oxidizing or reducing gas through the reactor at the initial temperature desired for effecting the reaction. It is preferred, however, to pass a stream of relatively inert gas, such as nitrogen, carbon dioxide, methane or other saturated hydrocarbon, through the reactor initially to fiuidize the contact material and purge the reactor of undesirable gases. Thereafter. the introduction of a. stream of the oxidizing or reducing gas, as the case may be, is initiatedat thedesired temperature. Alternatively,

the passage of the oxidizin or reducing gas stream may be initiated at a relatively low temperature after which the temperature of the gas stream is gradually raised to the necessary oxidizing or reducing temperature. Both the oxidation and reduction treatment are preferably continued, by the passage of the treating gas through the reaction zone at a velocity effective to produce the desired fluidized condition of i the finely divided contact material. until oxidation or reduction, as the case may be, is substantially complete. Completion of oxidation is indicated by the increase in. oxygen content of the eiiiuent gas and completion of reduction is indicated by the substantial absence of water in the gas stream emerging from the reduction reaction. In some instances partial reduction of the catalyst mass comprising iron is sufiicient to produce the desired catalytic eifect during the hydrogenation of carbon monoxide. An iron catalyst comprising less than 50 weight per cent elementary iron disregarding the presence of promoters and supports is within the scope of this invention.

After reduction of the catalytic material and in order to activate that material to its maximum activity for the hydrogenation of carbon monoxide in accordance with this invention, the catalytic material is subjected to an activation treatment which comprises gradually, and preferably progressively, raising the pressure from below 85 pounds per square inch gage to above 100 pounds per square inch gage to the desired operating pressure during contact of at least 200. standard cubic feet of synthesis gas having a concentration of at least about volume per cent carbon monoxide per hour per pound of catalyst calculated as the elementary metal and disregarding the presence of promoters and supports.

In regenerating and activating the catalyst which has been used in a finely divided condition for the synthesis of organic compounds by the hydrogenation of carbon monoxide, the catalyst may be continuously or intermittently withdrawn from the synthesis reaction zone and subjected to successive treatments of oxidation, reduction and activation in a single or a plurality of zones. Alternatively, the entire catalyst mass in the synthesis reaction zone itse f may be subjected to successive treatments of ox dation, reduction and activation, which method results in an intermittent synthesis process, whereas the former method permits a continuous synthesis process.

As used in this specification and claims, suspending the catalyst in a "fluidized condition" or by the fluidized technique has reference to the catalyst either when it is in the pseudo-liquid dense phase or when it is entrained and circulates in a continuous phase through the reaction zone. Fluid-bed technique has reference to the pseudo-liquid dense phase type of operation. The term regeneration has reference to treatment of a spent or partially spent catalyst by either oxidation or reduction or both. 0n the other hand, activation has reference to the special treatment of the catalyst according to this invention comprising treating either a fresh or regenerated catalyst to impart to the catalyst its maximumactivity for the production of normally liquid organic compounds without excessive formation of wax or relatively high-boiling suring adequate fluidization of a finely divided catalyst.

The invention in various modifications will be described further by reference to the accompanying drawings in which Fig. 1 is a View in elevation, partly in cross-section, of a relatively small reactor suitable for carrying cut themvention, and in which Figs. 2 and 3 are diagrammatic views. in elevation of arrangements of apparatus adapted for carrying out the new process in a continuous manner. This invention will also be described in greater detail by reference to specific examples of the improved process, which were carried out in the reactor shown in Fig. 1. It will be evident, however, that the arrangements of equipment indicated in the drawings merely illustrate certain modifications of the invention, which is capable of other modifications as will be evident to those skilled in the art.

In such further description as well as in the prior description pressures are expressed as pounds per square inch gage and volumes of gas as (standard) cubic feet measured at 70 F. and atmospheric pressure.

Referring to Fig. 1 reactor il consists of a length of extra heavy standard 2-inch steel pipe which is about 153 inches long and has inside and outside diameters of 1.94 inches and 2.38 inches, respectively. Reactor II is connected, by conical section l2, to an inlet pipe l3 made of extra heavy standard half-inch steel pipe having an inside diameter of 0.55 inch. Reactor ii is connected at the top, by means of conical section 14, with an enlarged conduit l5 comprising a length of 6-inch extra heavy standard steel pipe having an inside diameter of 5.76 inches. Conical section II and conduit [5 constitute an enlarged extension of reactor H which facilitates disengagement of catalyst from the gas stream after passage of the latter through a dense catalyst phase.

Conduit I5 is connected by means of manifold it with conduits l1 and 18 which comprise other sections of extra heavy 6-inch standard steel pipe. Conduits l1 and I8 contain filters 19 and 20 which are constructed of porous ceramic material which is permeable to the gas and vapors emerging from the reaction zone but impermeable to the catalyst particles carried by entrainment in the gas stream. Filters [9 and 20 are cylindrical in shape and closed at the bottom ends. They are dimensioned in relation to conduits l1 and [8 to provide a substantial annular space between the filter and the inner wall of the enclosing conduit for the passage of gases and vapors and entrained catalysts upwardly about the outer surface of thefilter. The upper ends of filters l9 and 20 are mounted in closure means 2| and 22 in a manner whereby the gases and vapors must pass through either filter l9 or filter 20 to reach exit pipes 23 and 24. Each of filters l9 and 20 is approximately 36 inches long and 4 inches in outside diameter, the ceramic filter walls being approximately of an'inch thick.

The greater part of reactor II is enclosed in a jacket 25 which extends from a point near the top of the reactor to a point sufliciently low to enclose the 3 inch length of conical section i2 and approximately 5 inches of pipe l3. Jacket 25 comprises a length of extra heavy 4- inch standard steel pipe having an inside diameter of 3.83 inches. The ends of jacket 25 1 are formed by closing the ends of the 4-inch pipe in any suitable manner, as shown, and sealed by welding. Access to the interior of jacket 25 is provided by an opening 26 in the top thereof through a 2-inch steel pipe. Jacket 25 is adapted to contain a body of liquid for temperature control purposes, such as water, or Dowtherm. The vapors which are evolved by the heat of reaction in reactor II are withdrawn through conduit 26, condensed by means not shown, and returned through conduit 26 to the body of temperature control fluid in jacket 25. The temperature control fluid in jacket 25 is maintained under a pressure at which the liquid boils at the temperature desired in Jacket 25. Electrical heating means (not shown) is provided in connection with jacket 25 to heat the temperature control fluid therein to any desired temperature, for use particularly when starting up the hydrogenation reaction.

In order to show all the essential parts of the reactor and associated catalyst separation means on a single sheet, a large proportion of the apparatus has been eliminated by the breaks at 21 and 28. For a clear understanding of the relative proportions of the apparatus reference may be had to the over-all length of the apparatus, from the bottom of jacket 25 to exit pipes 23 and 24, which is about 224 inches. In each of breaks 21 and 28 the portion of the apparatus eliminated is identical with that portion shown immediately above and below each break.

In the operations carried out in the apparatus of the drawing the catalyst recovery means comprising filters l9 and 20 is effective to separate substantially completely entrained catalyst from the outgoing stream of gases and vapors. The disengagement of solids from the gas stream is promoted by the lowered velocity of the gas stream in conduits l5 and remaining solids are separated on the outer surfaces of filters l9 and 20. The latter are employed alternately during the operation so that the stream of gases and vapors and entrained solids passes from conduit I 5 through either the left or right branches of manifold 16 into either conduit I! or conduit l8. During the alternate periods the filter which is not in use is subjected to a back pressure of gas which is introduced at a rate sufiicient to dis odge catalyst which has accumulated on the outer surface of the filter during the active period. Such "blowback gas and dis odged catalyst flow downwardly in the conduit enc osing the filter and into manifold IS in which the blowback gas is combined with the reaction mixture flowing u wardly from conduit IS. The great r part of the catalyst thus dislodged sett es downwardly into the reactor and is thus returned for further use. The blowback gas conveniently comprises recycle as such as from conduit ll In the operation of the apparatus of the drawing the desired quantity of powdered catalyst is f -introduced directly into the reactor through a suitable connection. not shown. in conduit I5. After any desired preliminary activation treatment the temperature of the fluid in jacket 25 is adjusted. by the heatin means mentioned above and by the pressure control means. to the temperature desired to be maintain d in jacket 25 during the reaction. After the catalyst mass has reached the reaction temperature .the introduction of the reaction mixture through pipe I3 is initiated. The reaction mixture may be preheated by means not shown approximately to the reaction temperature prior to its introduction through pipe l3 or the reactants may be heated to the reaction temperature through the passage thereof through that portion of pipe l3 which is enclosed by jacket 25 and by contact with the hot catalyst. It will be understood, furthermora'that the enclosure of pipe l3 in jacket 25 is not necessary to the invention and that the reactants may be heated to the reaction temperature solely by contact with hot catalyst.

Pipe I3 is dimensioned with respect to reactor II and the desired superficial velocity whereby the linear velocity of the gases passing through pipe I3 is sufilciently high to prevent the passage of solids downwardly into pipe l3 against the incoming gas stream. A ball check valve, not. shown. is provided to prevent solids from passing downwardly out of the reactor when the gas stream is not being introduced into pipe I3.

The reaction eflluent from reactor l I'is removed therefrom through either or both conduits 23 and 24 and passed by means of conduit 3| to a primary condensation unit 32. Condensation unit 32 comprises a jacketed accumulator in which steam is passed around the accumulator through a jacket to cool the reaction eflluent to a temperature of about 300 F. at the operating pressure existing in reactor ll. Cooling of the reaction eiiluent at the operating pressure to about 300 F. condenses the relatively high-boilin hydrocarbons and waxes which are removed from the condensation unit 32 through conduit 33. Uncondensing vapors are removed from condensation unit 32 and passed through a condenser 36 to accumulator 31. Condenser 36 cools the reaction eiiluent to a temperature below about F. and results in the accumulation of two liquid phases in accumulator 31. The two liquid phases formed in accumulator 31 comprise a heavy water-rich phase containing dissolved oxygenated organic compounds and a lighter hydrocarbon-rich phase which also may contain some oxygenated organic compounds having more than four carbon atoms per molecule. The two liquid phases are withdrawn from accumulator 31 through conduit 38 for subsequent recovery and purification by means not shown, such as by distillation and extraction. Uncondensed components of the reaction eilluent comprising unreacted hydrogen and/or carbon monoxide, methane and carbon dioxide are removed from accumulator 3'! through conduit 39. These gases may be vented to the atmosphere if desired, or may be recycled through conduit 4| to inlet conduit l3 of reactor II to supplement the feed thereto and to alter the ratio of hydrogen to carbon monoxide in reactor II. The presence of carbon dioxide and methane and other dfluents in the recycle stream serves to strip the relatively organic compounds and waxes from the catalyst particles in reactor II and is thus an aid in preventing settling of the fluid-bed of catalyst.

The following examples are illustrative of the procedure for starting-up a process for the hydrogenation of carbon monoxide with a freshly reduced catalyst comprising iron. Since the ex-- amples are illustrative only of the starting-up procedure and in some cases the actual operating conditions for producing the desired product, they should not be considered unnecessarily limiting and are offered merely as a better understanding of the improved process of the presentinvention.

The runs illustrated in the examples were carried out in apparatus substantially the same as that shown in Fig. 1 of the drawings. The results of each operating run are reported in conwith water to remove potassium oxide.

13 ventional tabular form. The contraction and yield of oil and water may be taken as indications of catalyst activity. The yield of oil repreorganic compounds since it does not include most of the gaseous hydrocarbons made or the organic compounds soluble in the water product. The yield of water represents the aqueous layer recovered in the primary and secondary receivers and includes organic compounds soluble therein.

The contraction in per cent is a measure of the amount of conversion of carbon monoxide.

The break-down of the reaction products showing the distribution of particular products obtained by CO conversion is an indication of the selectivity of the catalyst.

EXAMPLE I Catalyst for use in this operation was prepared by suitable treatment of a mixture of iron oxide and alumina and potassium oxide, previously prepared by fusion of the alumina and potassium oxide in molten iron oxide, for use in catalyzing ammonia synthesis. This material consisted principally of iron oxides and contai ed about 2.7 per cent alumina, 1.7 per cent potass um oxide,

04 per cent silica, 0.7 per cent titania and about 94 per cent iron oxides. It was first ground to a 6-20 mesh size and then subjected to leaching This treatment reduced the content of potassium oxide from 1.7 per cent to 0.5-0.6 per cent. The treated granular material was dried at about 200 F. overnight and then reduced in a stream of hydrogen. 7

In the reduction treatment a heated stream of hydrogen was passed through the granular mass, treat d by heat exchange with tap water to remove most of the water formed by the reduction reaction, and then recirculated. Reduction was initiated at about 700 F. The temperature of the catalyst mass was then raised to about 1350" F. in four hours, while continuing the flow of the hydrogen stream. This condition was maintained for two hours longer, during which time the reduction was substantially completed, as evidenced by the practical cessation of water formation. The reduced mass was then cooled to room temperature in the hydrogen atmosphere. Partial reduction of the catalyst, for example where only about 50 per cent or less of the iron oxides are reduc d to Fe, is within the scope of this invention; however, substantially complete reduction is preferred.

The reduced catalyst was then ground, first in a hand mill and then in a ball mill. to the desired degree of fineness. Throughout this period the catalyst was not; permitted to come in contact with air, the gr nding operations being conducted in an atmosphere of C02. The catalyst powder contained .about 93 weight per cent iron (Fe) and less than about 0.7 weight per cent potassium calculated as the oxide and had the following screen and Roller analyses:

Table I ROLLER ANALYSIS scamcN ANALYSIS U. 8. Std. Sieve:

+40 mesh Trace 40-60 Trace -80 Trace -100 Trace l00120 Trace -140 Trace -200 per cent" 122 ZOO-Pan do 86.3

About 20. pounds of the catalyst thus prepared were charged into reactor ll through an inlet (not shown) in section I5 of Fig. 1 of the drawings. During this operation the catalyst was maintained in an atmosphere of carbon dioxide and a small stream of 1 or 2 cu. ft. per hour of carbon dioxide was passed upwardly through reactor I to prevent packing of the catalyst. After the catalyst was charged to the reactor the carbon dioxide stream was replaced with a stream of hydrogen which was passed upwardly through reactor H at the rate of 10 to 20 cu. ft. per hour. The outlet pressure on the reactor was then raised to about 200 pounds and the temperature in the reactor was raised to approximately 450 F. by means of the heating coils around jacket 25. At that point the hydrogen flow rate was increased to 50 cu. ft. per hour and the temperature was then raised to about 550 F. The above flow rate is equivalent to a superficial linear velocity of about 1.3 feet per second in reactor H which produced a pseudo-liquid dense phase of fresh catalyst having a height of about 100 inches.

After reactor II had reached a temperature of about 550 F., approximately cubic feet per hour of fresh synthesis gas having a mol ratio of hydrogen to carbon monoxide of about 2:1 and having less than about 10 volume per cent components other than the reactants was substituted for hydrogen and passed through reactor I I under the above conditions. Approximately 715 cubic feet per hour of recycle gas from the secondary condenser or accumulator 31 was recycled to inlet 13 of reactor II, which increased the ratio of hydrogen to carbon monoxide and the amount of other components in the feed. The pressure was raised to approximately 250 pounds per square inch gage after about one hour operation. After the first nine hours of operation it was apparent that the catalyst was no longer fluidized. Synthesis gas rates were reduced and the recycle gas rates were increased in order to raise the hydrogen to carbon monoxide ratio in reactor I I in an attempt to get the reactor catalyst fluidized again. This operation was unsuccessful in fluidizing the catalyst. After approximately 32 hours operation the unit was disassembled in order to remove the catalyst therefrom. The catalyst had adhered to the walls of the reactor and it was necessary to chop out a large portion of the catalyst and clean the walls with a wire brush. The catalyst was present in large agglomerates and chunks and before reuse had to be reground or repulverized and washed with xylene to remove wax. The walls of the reactor were also washed with xylene to remove the wax formed during the rum. Table II below shows the operating data and results. for the run initiated at 200 pounds per square inch gage andat a temperature of about 550 F. It will be noted that although the contraction was fairly high, the production of oil was only fair and a relatively high proportion of wax was formed. ually increased until an average contraction of Table H J about 65 per cent was obtained. Recycle gas was introduced with the fresh feed at the beginning I 3 3 2 q of the 150 pounds per square inch gage run in ggfi% g3 g 14 order to raise the inlet feed rate to above 500 Operating Conditions: cubic feet per hour. Within a few hours after ggiggg g- 550 570 560 550 560 550 the operating pressure of 200 pounds per square Pressure, p. s. 1. guge- 200 200 250 250 .inch gage was reached, it was apparent that the fir e ir fi ri r llTIlIlI'. 200 11 s 1 e a catalyst was no longer fluidized but had settled. oas nat c s ogglr 260 170 85 m 50 The mol ratio of total feed to fresh feed at this Rg uef I II II 715 400 m5 580 m 4 0 time was approximately 2:1. Several hours of (Tim) 865 700 665 operation were spent trying to recover the fluid- Contraction, per cent... g2 g5 ized condition of the catalyst in the reactor with- OiLccJmfifrosh feed 34 4 wax gmwmmeshreed 4 4 4 PI rgg ed out success. Table IV below shows the operating Water, cox/m. fresh feed 97 97 97 conditions and results obtained in the run described in Example II. It should be noted here The accumulation of wax on the powdered .also that aconsiderable amount of wax was catalyst throughout the run illustrated in Table formed during the operation. When the unit was IIis shown below in Table III. It should be noted shut down the catalyst had to be scraped from that the production of wax was excessive during the walls of the reactor.

. Table IV Hours on Condition l l 2 l l l l l G 2 2 Catalyst Age,H'ours...- l 2 4 5 6 7 8 9 l5 17 19 Operating Conditions:

Temperature F.- I 4 Av. Catayst... 440 440 445 440 440 445 455 520 550 535 Max. Catalyst.. 452 455 455 445 445 453 451 550 514 554 Byn.lecd gas... 570 545 510 490 490 495 490 510 570 590 Pressure, p. s. i. gage- Reactor outlet o 15 so 75 100 150 200 250 250 100 Jacket 0- -0 0 Gas Rates, SCF/H- Synthesis Gas 35 100 120 200 250 215 245 270 180 Recycle 120 180 290 330 15c Inlet (total) a5 55 100 120 200 250 335 425 550 510 225 Total Cu. Ft./#cat 4.5 14.5 20.5 30.5 43 54 15 242 292 314 Results:

Contraction, per cent 35 65 17 Oil, calm. fr. feed 9 60 60 Wax, g1n./m. fr. feed 2 8 8 Water, cc./1n. ir. feed 26- 117 117 the operation which resulted in the agglomeration of the catalyst and settling of the fluid-bed in reactor I I.

In this example a similar catalyst prepared in a similar manner and having similar characteristics'to the catalyst of Example I, was brought up to a starting temperature of about 425 F. in the manner described in Example I. In this example, however, the pressure was not increased but was about atmospheric pressure atthe start of the run when fresh synthesis gas having a mol ratio of hydrogen to carbon monoxide of about 2:1 was introduced into the fluid bed of reactor H. Both the pressure and temperature were progressively increased during this run of Example II. The time required for increasing the pressure from atmospheric to about 100 pounds per square inch gage was equivalent to approximately 43 cubic feet of synthesis gas per pound of catalyst and the time required for increasing the pressure to 250 pounds per square inch gage was equivalent to'about cubic feet per pound of catalyst. The temperature was maintained sub,-

stantially constant during the procedure of rais- 1 ing the pressure to 250 pounds per square inch without recycle.

It was concluded in Example I that the initial activation of the catalyst was not proper and that the bed tended to settle because contact be-- tween catalyst and synthesis gas was effected at too high an initial pressure. In Example II it was concluded that the bed agglomerated and settled because the pressure was raised too rapidly to the desired operating pressure of 100 or 150 pounds per square inch gage.

' EXAIVIPLE III In Example III a similar catalyst and a similar procedure for initiating the reaction was carried out as in Examples I and II except that the initial pressure reaction was at pounds per square inch gage and the pressure was raised above pounds per square inch gage. The period of time required for raising the pressure of '85 pounds per square inch gage to above 100 pounds per square inch gage was equivalent to more than 500 cubic feet of synthesis gas per pound of catalyst. This operation was a single-pass operation The initial catalyst temperature was about 500 F. when the synthesis gas was first introduced and was gradually raised to 650 F. during the succeeding 50 hours of operation.

After 45 hours of operation or after approximately 500 cubic feet of synthesis gas per pound of catalyst had been introduced, the reactor pressure was gradually raised to pounds per square inch gage. At 140 pounds per square inch gage the catalyst exhibited high activity and the bed was maintained in a fluidized condition at all times. No difficulty was experienced in effecting 5mm and mm m mm wmmmm Mmumwm a a can m m mm m m n m m m c m mmmm lawm h a. m mamas.

saga Mn W t tmmmomm mmmmm a mam mmmmmmw d e mwmmm a mmmmmmm mmmme m m cg c d l m u m ma m m mwmam oka m ma mmmwm awmmwmm with t but arent n and n the catalyst carbo sity of the catalyst. rated density of t a linear velocity After about 600 density of the catalyst Table W! d after it has built up to an app ibrium. The accumulation of e fresh catalyst had an ac EXAMPLE IV Free-- It appeared that the carbon formation freshly reduced catalyst was rapid at firs cease relatively heavy hydrocarbon upo considerably changed the den Th about 80 pounds per cubic toot a or about 0.5 foot per second.

hours of operation the utlon, Wt. per cent:

was about 50 pounds per cubic foot.

or: on Btrca lter Ha Trea Anal t War Wax and Fe Iron Oxid I100 lbs. Fe at/lOO lbs.

an Kit. In'o.

Iron Distrlb Oil,

Total B0 Hours A Oatal Lbs. 0 Lbs. C

p DI procedure of Ex- Table V was ma m mam a m mam m an and man a as mam m mm a a me a an r... mam mm a. a was a... mm at mam m mm as aw ma aw a yst was mainr over 1200 adperation but dropped 25 quil mount after about 20 the successful starting-up ample III are shown in Table V below.

F.- Av. Gatal Total Cu. Ft Results:

The operating conditions an Temperature,

there any difilculty in fluidizin The yield of wax during the Example III were negligible. 0f 3 cos. per cubic meter 0 observed at the start of the 011 rapidly to a negligible a hours of operation.

for

Hours on Condition.

Catalyst Age, Hours.

Operating Conditions:

Contraction Wax, gin/m. ir. feed.-.

1 Hydrogen added to synthesis gas.

Following the initial run or startin cedure with the catalyst the catal A wide range of operating varitained in continuous operation to ditional hours.

1 pressure in accordance with the procedure outlined in E tact between catalyst and synthesis gas having a hydrogen to ables were investigated at pressures up to 250 A catalyst was prepared as in Example I and pounds per square inch gage and synthesis gas was brought up to temperature an ratios of 1.7:1 to about 3.5:1 and recycle gas ratios of total feed to fresh feed up to about 2.8:1. xample I. In this example, however, the start- These additional hours of successful operation at ing pressure at which the initial con Table VI Av. Oat. Temp., F...

Pressure, p. s. i. gage Space Velocity, SCF/Hr./Lb. Fe

Recycle Ratio i. Synthesis Feed Gas:

Tot. Inlet rad, H, c0-. Observed Yields: I

Contraction, Vol. per cent.

CO Conversion Water, ec./m.

1 Based on fresh iced.

I omitted/mob (and.

varying conditions are shown in Table VI below.

Catalyst Age, Hours Operating Conditions:

carbon monoxide ratio of. about 2:1 and having about volmne per cent of components other than the reactants was about 15 pounds per square inch gage. The temperature of the catalyst upon initial contact between catalyst and synthesis gas was about 500 F., and the maximum catalyst temperature during treatment was about 630 F. The pressure was gradually raised to 100 pounds per square inch gage over a period of time equivalent to a contact of over 200 cubic feet of synthesis gas per pound of catalyst. The operation was a singlepass operation without recycle. After full activity of the catalyst had been exhibited at successful results obtained on this particular catalyst after pretreatment for pressures varying from 50 to 150 pounds per square inch gage are shown in Table 11: below.

Table 12:

A B 0 D E E G H C t ly A 9, Hours 214 231 257 455 491 008 662 787 Operating onditions:

Temperature, 510 550 550 519 550 510 955 610 Pressure, 9.5.1. case-.. so so so so 150 50 50 so gn Velocity. 1, 055 1, 200 1, 320 1,215 1, 855 1, 195 840 915 R 2 Gas ter reactor 1. 9 1. 9 1. 9 2.0 2.0 2.1 1. 9 1. 4

Contraction. went 48.9 50.1 41.2 50.0 52.9 24.6 51.3 49.9 Perce 0 v ion 31.1 86.9 15.4 99.5 93.1 41.5 94.0 91.3 v(JO Distribution, Peroent- C0 00 39.0 34.2 31.9 30.8 21.9 36.8 31.3 12.2 9.2 10.1 10.9 11.1 22.8 10.4 00 0 ted o d 3. 2 4 3 2 2 3'3 1 --0' xygena 0m 1111 S Yields, calm. Fresh feed: p0

s2 s5 23 a4 21 v25 44 04 s 5 14 12 14 5 1s s e 2 11 1 5 11 Light Naphtha 15 19 23 13 19 27 68v 42 50 42 41 5a 13 1s 49 Oxygenated Cmpds in water.. 7 7 5 6 10 2 5 3 star. 58 52 42 01 19 21 14 59 about 100 pounds per square inch gage the pres- EXAMPLE V sure was decreased to about 80 pounds per square inch gage. At the lower pressure activity of the catalyst was maintained as a result of the starting-up procedure of Example IV. Throughout the starting-up process of Example IV continuous-and smooth operation was effected. No difficulty was encountered in maintaining the bed. in a fluidized condition. The operating conditions and the results obtained during the starting-up procedure of Example IV are shown in Table VIII below.

Table VIII Hours on Condition 3 5 5 33 12 Catalyst A e, Hours 3 8 13 46 58 Operating auditions:

, Temperature. F.

ya v60 90 130 145 Total G1 FJiOat 6 21 44 259 345 Results:

Contraction per cent 65 40 46 42 011, calm. esh feed 0 2 4 65 Wax, gm./m.' fresh feed- 0 0 0 6 6 Water, calm. fresh teed..- 66 106 50 40 It should be noted that the wax formation during the starting-up procedure was relatively low. It was concluded from Examples 111 and IV that it was necessary to initiate contact between catalyst and fresh feed gas at a pressure below about pounds per square inch gage and to increase the pressure gradually to about pounds per square inch gage over a period of time equivalent to the contact of at least about 200 cubic feet of synthesis gas per pound of catalyst. Preferably, the pressure should be gradually raised from the initial pressure to u it had agglomerated. Table X below illustrates Example V uses a somewhat different catalyst than the preceding examples. The main difference in the catalyst of Example V and the catalyst of the preceding examples is that the present catalyst contained a relatively high alkali content, in particular, a potassium content calculated as oxide of about 1.2 weight per cent after the catalyst was reduced. This catalyst was prepared in a similar manner as the catalyst in Example I and the procedure followed preparatory to initial contact with synthesis gas was also the same as that described in Example I. The initial contact between a synthesis gas comprising 2. mol ratio of hydrogen to carbon monoxide of about 2:1 was effected at a temperature of about 480 F. and at atmospheric pressure. The pressure was raised to about 100 pounds per square inch gage over a period of time equivalent to less than about 50 standard cubic feet of fresh feed per pound of catalyst. The pressure was maintained at 100 pounds per square inch gage for considerable length of time in orderto determine the activity of the catalyst and in an attempt to maintain the catalyst bed in a fluidized condition. Recycle gas was initiated after a pressure of 100 pounds per square inch gage had been reached in an attempt to retain the catalyst in a fluidized condition. Even at relatively high linear gas'velocities the catalyst tended to agglomerate and settle out. The yield of oil and water were relatively low during the operation and the wax yield was relatively high. After unsuccessful attempts had been made to obtain good conversion yields and to maintain the catalyst in a fluidized condition, the reactor was shut down. It was necessary to scrape a large part of the catalyst from the reactor wall where low shows the operating conditions and results obtained in Example VI in the unsuccessful attempt to bring about the maximum activity of the high alkali catalyst.

Table XI -Hoursoncnditlon 3 8 5 5 2 6 2 2 7 CatalystA .Hours 8 6 11 16 18 24 26 28 36 Operating onditions:

'Icm aturc F.-

1.01mi mwuommsmscomws MLLOBtBl 646 549 560 623 m 578 574 620 Synthesis eed 027 628 620 626 615 576 615 am 600 Pressure, p. s. i. gage- Reactor Outlet 0 i5 30 00 100 85 86 85 100 set ..l0 1010 9 0 0 9 9 GasRste,80F

g ynthesis as 50 85 ll U 135 140 220 215 230 Remus otsl O.F./l Cat 7 18 28 52 66 11B 130 152 222 CoutraetlcnVoLp ercent so 40' 36 85 35 30 37 0il,ec./m.'of Fres feed 8 8 8 8 8 8 8 9 Wax, gmJmJfresh feed 13 13 13 1.3 9 4 4 4 4 Watcr,cc./m.'freshfeed 1, 1 1 1 11 1c 1a 1a 21 EXAMPLE VII inadequate for starting-up the high alkali cata- In Example VII still another attempt was made Table X v 1 Hrs. on Condition l 3 2 2 3 3 3 1 2 6 4 2 1 Catalyst Age, Hrs l 4 6 8 11 14 l7 18 Z) 26 30 32 33 Operating Conditions:

em rature F- v.0ataiyst 480 500 500 600 610 550 685 560 540 600 550 540 560 Max. Cat 484 505 606 5% 677 614 694 680 795 672 656 065 By'n. Feed 660 670 865 636 660 570 670 580 690 530 545 no Pressure, p. s. i. gage- Reactor Outlet 0 16 30 60 100 100 100 100 100 100 100 100 100' Jae 0 0 0 0 0+9 16-11 0 0 0 0 0 0 0 Gas Rates ECF/H- Syn. as- 37 65 95 110 130 220 310 220 120 70 30 95 Recycle-. 86 130 350 B0 Inlet 215 350 600 110 'iotal Cu. Ft./lb.Cat- 2 7 13 22 38 70 120 153 176 211 225 M6 250 Contraction Per Cent l 8 9 60 50 45 70 81 0il,ec./m. r.feed. 0 0 0 0 20 20 f!) 12 l2 i2 44 44 44 WaX,gm./m.' infeed 0 0 0 0 8 8 8 4 4 4 8 8 8 Wster,cc./m.' oiiresh feed 0 O 0 0 40 40 40 22 22 22 71 71 .1

EXAMPLE VI with. the relatively high alkali catalyst to initiate A second attempt was made in Example VI to start-up with a high alkali catalyst under conditions to maintain maximum activity and the catalyst in a pseudo-liquid fluidized condition in the reactor. The catalyst and operating procedure were the same as those in Example V. The catalyst had an alkali content of potassium calculated as the oxide of about 1.2 weight per cent after reduction of the catalyst. In Example VI the catalyst was initially contacted with the synthesis gas in a single-pass operation at a temperature above about 500 F. and at substantially atmospheric pressure and thereafter the pressure was gradually raised to 100 pounds per square inch gage. In this operation the time required to raise the pressure to 100 pounds per square inch gage was equivalent to about cubic feet of synthesis gas per pound of catalyst. Because of the difllculty encountered in controlling the catalyst temperature, the pressure was dropped to 85 pounds per square inch gage for ten hours. Then the pressure was agai raised to 100 pounds per square inch gage. In this run the conversions were low as wellas the hydrocarbon yields. As the result of the agglomeration of the catalyst it was impossible to maintain the catalyst bed in a fluidized condition throughout the operation. After about 35 hours of operation the reactor was shut down and the catalyst scraped from the walls of the reactor. It was concluded from the operations of Examples V and VI that the starting-up procedures of these examples were inadequate. Table X! bethe hydrogenation reaction under conditions such that a fluid bed of catalyst could be maintained and under conditions such that maximum activity of the catalyst could be realized at operating conditions. This catalyst was similar to the catalyst of Examples V and VI and contained about 1.5 weight per cent potassium calculated as the oxide. The catalyst was prepared in the manner described in Example I and was brought to a temperature of about 500 F. and to a pressure of about 60 pounds per square inch gage before initiating contact between the catalyst and synthesis gas containing a 2:1 mol ratio of hydrogen to carbon monoxide. The pressure was raised gradually from about 60 pounds per square inch gage to about 100 pounds per square inch gage during contact with more than 500 cubic feet of synthesis gas per pound of catalyst. Simultaneously with the rise in pressure, the temperature was raised from about 550 F. to above about 600 F. After the pressure had reached 100 pounds per square inch gage the pressure was further raised gradually to 250 pounds per square inch gage, the predetermined operating pressure. vIt was apparent after the pressure had been raised slowly to 100 pounds per square inch gage that the catalyst had become highly active and throughout the remainder of the operation the fluidized bed was maintained without difilculty. All of the runs of Example VII were recycle runs.

The operating conditions and results obtained by this method of starting-up a high alkali catalyst are shown in Table XII below. It will be 23 noted that the catalyst exhibited high activity throughout the runs made.

pressure during contact with at least about 250 cubic feet of synthetic gas per pound of catalyst.

1 Based on fresh food.

It was concluded in Example VII that, in order to obtain a high alkali catalyst at maximum activity without danger of improper fiuidization of the catalyst bed, the catalyst and synthetic gas must be initially contacted at a, pressurev below about 85 pounds per square inch gage and raised gradually to the desired operating Table XII Fig. 2 of the drawings is a diagrammatic illus- 3 s a a 2 25 a s4 3 2 12 21 2a 4a 61 85 560 665 510 510 510 e10 e 620 564 516 592 see 525 658 622 640 460 480 410 480 500 545 615 510 so s5 85 so so 85 85 ran Jacket 1s 19 19 1s 2a 30 22 Gas Bates BCF/H- Syn. das (fresh iced)-.. 55 100 as as 115 150 Recycle 110 145 160 160 160 145 140 240 245 225 225 235 200 290 136 231 200 ass 591 1,087

Contraction Vol. Per 0ent..- 65 40 30 40 45 68 0il,-cc/m.'f'mshieed. 13 2 9 1o 13 a 2 12s 2 s 12 5 4 s s so as 101 5 11 144 84 After the catalyst had been fully activated-in tration in elevation of one particular type of conaccordance with this procedure, tests were made tinuous process for the synthesis of hydrocarbons upon the catalyst at other conditions of temperain which a contact mass of iron catalyst is subure and pressure to determine its continued active jected to successive treatments of oxidation, reity and ability to be maintained in a fluidized 25 duction, activation and synthesis with the use of condition. Pressures varied from 120 to 250 four zones whereby the synthesis reaction may pounds per square inch gage while good yields of be effected in at least one of the zones at all oil were exhibited throughout the varying conditimes. In accordance with this modification, tions of pressure. Particularly, the highalkali assume that reaction chamber II is being opercatalyst exhibited high yields of oxygenated 30 ated on the synthesis of organic compounds by organic compounds which was not apparent with the hydrogenation of carbon monoxide in the the low alkali catalyst even after the proper actipresence of a finely divided iron catalyst in a vation treatment. The oxygenated organic comfluidized condition. The other zones I2, 13 and pounds recovered from the water and oil phase. I4 are being operated on oxidation, reduction from the condensers comprise acids, ketones, as and activation, respectively. In this method of aldehydesand alcohols. For a more complete operation all four zones are simultaneously being discussion of the products produced with the high operated on different reactions; thus, in zone ll alkali catalyst, reference may be made to my the synthesis reaction is being efiected by passco-pending application Serial No. 725,835, filed ing a gaseous mixture of hydrogen and carbon February 1, 1947 in which a detailed discussion monoxide thereto through conduits I8 and 82 by the Prodlwts Produced w the high alkali adjustment of the proper valves and from which catalyst is presented. Typical results with the zone the reaction eflluentis removed through eon Present activated high a al catalyst is shown in duit 83. Simultaneously, the catalyst in reaction Table XIII below. zone 12 is undergoing an oxidation treatment in Table x, accordance with the previous discussion and this oxidation treatment is effected by passing oxy... A B C D E F gen thereto through conduits I6 and 84 and by adjustment of the proper flow valves. Flue gases are removed from zone 12 through conduit 88. 31.35; 3533:333 24 24 24 24 Simultaneously, a catalyst mass substantially av. at aiyst 'lgengg F- g g"; 925 free from carbonaceous deposits in reaction zone c5221 i'l l'ccityf r'cI f I3 is undergoing reduction by passing a reducing g -k -g g gas, such as hydrogen, therethrough by means 19;211:111 2.68 3.02 ".05 2.12 2.00 1.01 of conduits l1 and SI, and adjustment of the 111 M01 proper valves and regeneration gas is removed F F .4 1.6 .5 1.0 1J1??? T321 22 ii ii 2.0 1 .0 2.2 from zone 13 through conduit 93. Simultane- Contraction v01 Per ously, a regenerated or freshly reducedcatalyst ocm..- 58.9 65.4 5.5.1 43.1 50.1 55.9 mass is bein subjected to an activation treatcng jjf f ilfl fi 99.3 94.4 88.7 85.3 86.5 mo ment in accordance with the teaching of this oii cc/irl 01 fresh fecd.. 116 121 91 11 as 100 an invention in reaction zone 14 and this activation fga f f fff f ff fff 77 56 69 so 38 54 treatment is accomplished by passing a gaseous Water, ca m. of fresh mixture of hydrogen and carbon monoxide fiff 89 81 57 through reaction zone It by means of conduit CO- C0,P c a 22 23 COQOHH-L 15 em 2; 1; a 2; 13 u 2181:1161 removing an emuent therefrom from con (JO-10x3. Comp, Per u Cant 18 15 20 25 11 13 After the catalyst in the synthesis reaction zone I H has become substantially deactivated, and the carbonaceous deposits thereon. Accordingly. oxygen is passed through conduits It and I! to reactor II; hydrogen is passed through conduits II and It to reactor I2; and hydrogen and carbon monoxide are passed through conduits I8.and 92 and .conduit II to reactors I3 and I4, respectively. Conduits 82 and II are used when reactors II and I2 are on synthesis or activation. Conduit 8| is used when reactor II is on reduc tion. Conduit I! is used when reactor I2 is on either synthesis or activation, and conduit 89 is used when reactor 13 is on oxidation.

In this manner the continuous production of organic compounds by the hydrogenation of carbon monoxide in the presence of an active catalyst may be maintained. In some instances it may be unnecessary to subject the spent catalyst to oxidation as the removal of carbonaceous deposits may be conveniently accomplished during the treatment of the catalyst with hydrogen. Under such circumstances only three zones may be necessary and one of the aforesaid zones may be omitted. It may often be desirable to purge each of the zones before and after the various treatments are carried out in each of the zones. In this modification each of the zones may be purged with a substantially inactive gas, such as methane, nitrogen, carbon dioxide, or steam. The purging treatment is for a short period of time and substantially completely removes the reaction gases from each of the zones prior to the introduction of the next reaction gas mixture therein for the particular treatment to be undertaken in that zone.

Each of the zones I I, I2, I3 and I4 goes through a typical cycle involving the following steps shown in Table XIV below.

Table XIV Activation Synthesis Purge Oxidation (may be omitted) Purge Reduction Activation In the foregoing detailed description of the invention reference has been made to the operation of the type exemplified by Figs. 1 and 2 in which a starting-up procedure and activation of the catalyst is carried out in the synthesis reactor itself. The invention includes within its scope the operation of a single synthesis reactor in a continuous manner with continuous or intermittent removal of a portion of the catalyst mass from the reactor for transfer to separate regeneration and activation chambers. This modification of the invention is illustrated by Fig. 3 of the drawings which is a view in elevation diagrammatically illustrating an arrangement of apparatus adapted to maintain the synthesis reactor in continuous operation with simultaneous regeneration and activation of spent catalyst. The apparatus of Fig. 3 comprises a synthesis reactor III, an oxidation unit I31, a reduction unit I46, and an activation unit I54, suitably connected for carrying out intermittent or continuous regeneration and activation of a portion of the catalyst mass from synthesis reactor III.

Synthesis reactor III is adapted to contain a mass of finely divided catalyst to be maintained in a fluid-bed condition with an upper level or interface at about the height indicated by numera1 I22. Synthesis reactor III is also provided with a suitable cooling means III for removing heat from recycled catalyst in order to maintain a temperature within reactor III at the desired temperature level. Cooling means III suitably comprises a plurality of tubes containing a vaporizable liquid and connected to a common header which is maintained at the pressure necessary to cause the liquid in the tubes to boil at the temperature necessary to provide the desired rate of heat exchange with the circulating catalyst passing therethroughr Instead of the specific temperature control means shown as element III, any other conventional suitable temperature control means may be employed. such as indirect cooling tubes positioned within synthesis reactor I II, without departing from the scope of this invention.

The reaction mixture comprising hydrogen and carbon monoxide within a mol ratio of about 1:1 and about 3:1 and having a carbon monoxide concentration of at least about 15 to 25% is introduced into the system of Fig. 3 through conduit II3 which may be provided with a conventional compression means or pump not shown. Conduit II3 connects with the bottom of reactor III after passing through cooler III and has a cross-sectional area substantially smaller than that of the reactor such that the introduction of the reaction mixture into reactor III provides a sufiiciently high linear gas velocity at the inlet to prevent the passage of catalyst down-ward out of reactor II I into conduit H3.

In accordance with the modification of Fig. 3, the gaseous reaction mixture of hydrogen and carbon monoxide is passed through conduit 3, cooler Ill and upward through reactor III in contact with a mass of finely divided flu dized catalyst under conditions such that the catalyst issuspended in a pseudo-liquid dense phase and an interface I22 between the dense phase and a relatively dilute phase in the upper portion of reactor III is formed. To assist in the disengagement of entrained catalyst particles in the dilute phase, a cyclone separator or ceramic filter (not shown) may be inserted within the upper portion of reactor III to separate finely divided entrained catalyst particles from the gaseous eiliuent. The gaseous effluent containing the products of the process, such as hydrocarbons and oxygenated organic compounds, and other components, such as steam and unreacted hydrogen and carbon monoxide, is removed from reactor II I through conduit I23. Alternatively or additionally to the catalyst separating means positioned within the shell of reactor II I, conduit I23 may contain similar separating means, such as cyclone separators, for the removal of any entrained catalyst from the gaseous eilluent. Entrained catalyst thus separated from conduit I23 is returned to reactor III by suitable conduits or standpipes not shown. The gaseous eilluent passes through conduit I23, cooler I24 to accumulator I28. Cooler I24 may comprise a single or a series of condensation units for cooling and condensing the reaction eiiiuent. Accumulator I26 collects the condensate from coo er I24. The condensate forms an upper liquid organic-rich phase and a lower liquid water-rich phase. The aqueous phase is withdrawn from accumulator I28 through conduit I29 for disposal or for treatment to remove valuable oxygenated organic products dissolved therein. The liquid organic phase is removed from accumulator I26 through conduit ISI for further treatment and separation of organic products of the process by such means as distillation and extraction (notshown) Uncondensed components of the effluent, 'such as hydrogen, carbon monoxide and methane are withdrawn from accumulator I26 through conduit I21 and are recycled to reactor III through conduits I21 and I28. The recycling of the uncondensed components, such as hydrogen, methane and carbon monoxide, in the desired quantity can be used to alter the ratio of hydrogen and carbon monoxide .in the reaction zone to a desired predetermined value and also can be used to lower the partial pressure of relatively heavy or high-boiling hydrocarbons in the reaction zone I I I, such that. substantially complete vaporization of such hydrocarbons is effected. In the latter'respect it may be very desirable to recycle a portion of the gases through conduit I21 to the upper portion of the. catalyst mass in order to remove relatively high-boiling hydro-- carbons from the catalyst particles by vaporization and prevent their accumulation thereon which would cause a tendency of the catalyst to agglomerate and the catalyst mass to settle in reactor III. The recycled gases may be used also as an aid in maintaining the temperature in reactor III at the desired level. In this respect the gases may be recycled at a temperature below about 100 F. through conduits I21 and I28 in the desired proportion to maintain the temperature in any portion of the reaction zone at a desired value. Although only two conduits are shown leading to reactor III at separate heights, a plurality of conduits at a plurality of heights may be employed without departing from the scope of this invention, leading directly into both the dense and dilute phases.

As previously discussed, in order to maintain the temperature of synthesis reactor III at the desired predetermined level, a portion of the catalyst mass is continuously or intermittently withdrawn from the dense phase bymeans of standpipe H8 and recycled through conduit I I3 and cooler I I1. It is desirable when'withdrawing catalyst from reactor III through conduit II8 to strip and/or treat the catalyst in conduit II8 by introducing hydrogen or a hydrogen-rich gas into standpipe I I8 through a plurality of conduits such as conduits H9 and I2I. The introduction oi. hydrogen or hydrogen-rich gas above the ratio of hydrogen to carbon monoxidein the feed gas in conduit II3 can be employed to increase the hydrogen and carbon monoxide ratio in reactor I I I to a desired value.

The feed composition of hydrogen and carbon monoxide flowing through conduit II3 may be diverted through conduit H2 and introduced at a plurality of points longitudinally along reactor III through conduits I I 4 and II 6, if desired. In this modification, the synthesis feed gas may be introduced into the lower portion of reactor III as well as through a plurality of points longitudinally along reactor III. The longitudinal introduction of the feed gas into reactor III serves to replenish the consumption of the reactants as the reactant gases progress upward through reactor III. Also introduction of fresh feed gas serves to strip relatively high-boilin products from the catalyst particles and aids in preventing their condensation by decreasing the partial pressure of the higher boiling products. In this respect the multiple introduction of the feed into reactor I I I is similar in some respects to a series of reaction zones in which fresh feed is introduced into each zone. Reactor III may comprise a series of separate reaction zones in I which the ieedis injected into each successive synthesis reactor III at a high level of activity,

a part of such catalyst mass is removed from reactor III, revivifled or regenerated and returned thereto. Such reviviflcation or regeneration is accomplished by continuous or intermittent withdrawal of a portion of the catalyst from reactor III by means of a standpipe I33 or other suitable means. Standpipe I33 is aerated with a suitable aeration gas, such as. hydrogen, recycle gas or nitrogen. The portion of the catalyst mass withdrawn through standpipe I33 is introduced into conduit I36 through which an oxidizing gas, such as oxygen or air, is continuously flowing. The introduction of the catalyst into the flowing oxidizing gas results in the suspension of the catalyst in the flowing gases and the passage thereof to oxidation unit I31 which may comprise a suitable cylindrical chamber in which a mass of finely divided catalyst to be regenerated is maintained in a fluidized condition. When the catalyst is maintained as a fluid-bed in oxidation unit I31, an interface is formed between a so-called pseudo-liquid dense phase in the lower portion and a relatively dilute phase in the upper portion of oxidation unit I31, designated as numeral I38. The velocity of the upward-flowing oxidizing gas in unit I31 is preferably between about 0.5 and about 5 or 6 it. per second to suspend the mass in a pseudoliquid condition. A temperature is maintained in reactor I31 above about 800 or 900 F. whereby the carbonaceous deposits and relatively heavy hydrocarbons deposited on the catalyst are burned. A pressure between atmospheric and the pressure existing in the synthesis reactor III is appropriate for the oxidation reaction. Preferably, a pressure of approximately atmospheric is employed in oxidation unit I31. Flue gas comprising combustion products oi! the oxidation reaction is withdrawn from oxidation unit I31 through conduit I39. Suitable separating means for removal of finely entrained catalyst particles may be positioned in oxidation unit I31 and addi tionally or alternatively in conduit I39. Such catalyst separating means may comprise a cyclone separator or a Cottrell precipitator. If it is necessary to preheat the mixture of catalyst and oxidizing gas, this may be accomplished by positioning a preheater, not shown, on conduit I38. Cooling means, not shown, may also be provided for maintaining the temperature at an allowable maximum in oxidation unit I31, such as by the use of indirect cooling coils positioned within reactor I31. The temperature may also be controlled by injecting a suitable inert gas, such as nitrogen, steam, or carbon dioxide, into conduit I31 or directly into oxidation unit I31. After sufiicient contact time has been allowed for the substantially complete combustion of carbonaceous deposits and relatively heavy hydrocarbons from the catalyst particles, catalyst is continuously or intermittently withdrawn from oxidation unit I31 by means of standpipe MI and passed to a reduction unit I46 through conduit I44. Standpipe Ill is aerated with a suitable gas such as air, oxygen, nitrogen, introduced therein through conduit I 42. It is often desirable in aerating Standpipe I to introduce a relatively inert purging gas, such as nitrogen, steam, or carbon dioxide to purge the catalyst of en trained oxidizing gas prior to introduction into 29 conduit I44. Fresh make-up catalyst may also be introduced into the system at this point in the process, when necessary. such as by means of conduit I42.

Hydrogen or other reducing gas is passed through conduit I44 which picks up by entrainment catalyst from conduit Ill, and the resulting mixture passes to a reduction unit I45 for the re-reduction of the oxidized catalyst to produce a relatively active catalyst for the hydrogenation of carbon monoxide insynthesis reactor III. In the event that the oxidation of the catalyst particles in oxidation unit I3! is undesirable or unnecessary, the catalyst particles may be passed directly from synthesis reactor I II and conduit I33 to conduit I44 by means of a conduit or standpipe I43. If the pressure in conduit I44 and reduction unit I48 is substantially higher than in either synthesis reactor III or oxidation unit I31, it may be necessary to employ several lock hoppers and connecting standpipes or a Fuller Kinyon pump to boost the catalyst to the desired pressure. I

Reducing gas passes upward through a mass of finely divided catalyst in reduction unit I45 at a velocity between about 0.5 and about 5 or 6 feet per second whereby the catalyst is maintained in a pseudo-liquid fluidized condition. In reduction unit I45 the pseudo-liquid dense phase forms an interface I 41 between that phase and a relative dilute phase in the upper portion of reduction unit I45. Excess reducing gas and products are withdrawn from reduction unit I48 through conduit I48. Suitable means as previously described in connection with units III and I38 is provided for the removal of finely divided entrained catalyst from gases. The reduction of the oxidized catalytic material is effected at atmospheric or super-atmospheric pressures as previously discussed. In eifecting the reduction at about atmospheric pressure, a

temperature between about 900 and about 1600 F. is employed. In the reduction of the catalytic material at super-atmospheric pressures of about 200 to about 500 pounds per square inch gage, temperatures between about 600 and about 800 F. are employed. Sufficient contact time between the catalyst mass and reducing gas is permitted for the substantially complete reduction of the catalytic material to metallic iron. It may be desirable in order to conserve hydrogen to recycle a portion of the eiliuent gas containing hydrogen from conduit I45 to conduit I44 by means not shown.

Reduced catalytic material is withdrawn from reduction unit I46 through a standpipe I49 and intermittently passed to conduit I53 through which a gaseous mixture of carbon monoxide and hydrogen in a mol ratio between about 1:1 and about 3:1 and having a carbon monoxide concentration in excess of about 15 per cent is flowing. standpipe I49 may be aerated with a suitable aeration gas, such as hydrogen, steam or nitrogen, introduced through conduit I5I. If desired, a portion of the efliuent gas containing unreacted hydrogen from conduit I48 may be introduced into conduit I53 through conduit I52 to increase the hydrogen to carbon monoxide ratio above that ratio of the gas introduced initially into conduit I53.

A mixture of suspended and freshly reduced catalytic material and gases comprising hydrogen and carbon monoxide are passed upward through activation unit I54 at a sufficient velocity to maintain the catalytic material preferably in a pseudo-liquid dense phase condition with the "temperature and 30 formation of an interface Ill between a dense phase and a dilute phase of catalyst. An eiiluent is withdrawn from activation unit I54 through conduit I51. Suitable means (not shown) for separating entrained finely divided catalyst from the eiliuent is provided in reactor I54 or in conduit I51, as previously described. The activation of the mass of catalyst in activation unit I54 is effected by initially employing a temperature above about 450 F. and a pressure below about pounds per square inch gage and subsequently gradually and progressively increasing both the pressure to those corresponding to synthesis reactor III. The pressure in activation unit I54 is raised from a pressure below about 85 pounds per square inch gage to a pressure of about pounds per square .inch gage over a period of time equivalent to at least about 200 standard cubic feet of gas per pound of catalyst. After the pressure reaches about 100 or pounds per square inch gage in activation chamber I54, the pressure is increased to the operating pressure, for example about 250 pounds, at a maximum rate of 50 pounds per 100 standard cubic feet of gas contacted.

If the pressure existing in reduction unit I 45 is higher than 85pounds per square inch gage, the pressure on the catalyst is reduced prior to introduction into activation unit I54. 0n the other hand, if the pressure in reduction unit I45 is below 80 pounds per square inch gage, such as about atmospheric pressure, the catalytic material may be passed directly into activation unit I54 without a substantial reduction in pressure. In the event the pressure existing in synthesis reactor III for the synthesis of organic compounds therein is below 100 pounds per square inch gage, the pressure in activation unit I54 after reaching about 100 or 150 pounds per square inch gage or higher is reduced to approximately the pressure existing in synthesis reactor III prior to introduction of the activated catalytic material therein. Regenerated and activated catalyst is intermittently withdrawn from activation unit I54 through standpipe I58 and introduced into conduit I13 for recirculation back into synthesis reactor III as a highly active catalytic material which is suitable for use in the manufacture of normally liquid organic compounds without excessive formation of waxes, etc.

When the gaseous effluent from conduit I51 contains unreacted hydrogen and carbon monoxide, all or a portion of the effluent may be recycled to conduit II3 by means of conduit I59 and IBI. Similarly the effluent from reduction unit I46 may be passed to conduit II3 through conduits I48, I52, and IBI.

In most instances, the eiliuent from reduction unit I45 and activation unit I54 is richer in hydrogen than the feed composition to synthesis reactor III and may be substantially free from carbon monoxide. According to one modification of the present invention, this hydrogen-rich gas mixture from either or both units I45 and I 54 is introduced into conduit I I3 as the sole source of gases therein, and passes through cooler ill to reactor I I I. A mixture of hydrogen and carbon monoxide in a mol ratio of about 2:1 is then fed to synthesis reactor I I I through conduits I I 2, H4 and H6. This gas mixture in conduit III may have a hydrogen to carbon monoxide ratio lower than that desirable in synthesis reactor II I but by the introduction of hydrogen-rich gas therein through conduit II3, the desired hydrogen to carbon monoxide ratio in reactor III is achieved. Furthermore, contacting the catalyst from conduit Ill with a high hydrogen content gasserves to at least partially regenerate and remove carbonaceous and heavy hydrocarbon deposits therefrom.

Various modifications and alterations of the process and arrangement of apparatus of this invention may become apparent to those skilled in the art without departing from the scope of this invention. The probable theory and specific examples are oflered as a means for better understanding the invention and are not to be con"- strued as unnecessarily limiting to the invention.

I claim:

1. The starting-up procedure for the hydrogenation of carbon monoxide in the presence of a finely divided fluidized contact material comprising fused and freshly reduced iron which comprises, prior to establishing the desired operating conditions initially contacting said contact material comprising. freshly reduced iron with a gaseous mixture having a mol ratio of hydrogen to carbon monoxide between about 1:1 and about 4:1 at a pressure of at least about butbelow I about 85 pounds per square inch gage and at a temperature above about 450 F.,. passing said gaseous mixture upward through a mass of said freshly reduced contact material at a velocity suflicient to maintain the contact material in a fluidized condition, gradually raising the pressure above about 100 pounds per square inch gage during contact with at least about 200 standard cubic feet of synthesis gas per pound of iron calculated as the elementary metal, during the initial contact at which the pressure does not exceed about 100 pounds per square inch gage establishing a temperature of at least about 600 F., and subsequently establishing the desired operating conditions for the synthesis of organic compounds in the presence of said finely-divided contact material in a fluidized condition.

2. The starting-up procedure for the hydrogenation of carbon monoxide in the presence of a finely divided fluidized contact material comprising freshly'reduced iron which comprises,

prior to establishing predetermined operating conditions passin a gaseous mixture comprising hydrogen and carbon monoxide through said contact material at a linear velocity sumcient to maintain the contact material in a fluidized condition at a pressure of at least about 15 but below about 90 pounds per square inch gage, raising the pressure above about 100 pounds per square inch gage during contact with at least 200 standard cubic feet of synthesis gas per pound of iron calculated as the elementary metal, during the initial contact at which the pressure does not exceed about 100 pounds per square inch gage establishing a temperature of at least about 600 F., and subsequently establishing the predetermined operating conditions for the synthesis of the desired organic compounds in the presence of said finely-divided contact material in a. fluidized condition.

3. The starting-up procedure for the hydrogenation of a carbonoxide in the presence of a finely divided fluidized contact material comprising freshly reduced iron which comprises, prior to establishing predetermined operating conditions contacting said freshly reduced contact material with a gaseous mixture comprising hydrogen and at least about 15 volume per cent carbon monoxide, passing said gaseous mixture through said finely divided contact material at a linear velocity suflicient to maintain the contact 32 initial pressure of at least about 15 but below about pounds per square inch gage, raising the pressure to at least about pounds per square inch gage, contacting at least 200 standard cubic feet of said gaseous mixture per pound of iron calculated as elementary metal and raising the temperature to at least about 600 F. during the period the pressure is raised from the initial pressure to about 100 pounds per square inch gage, and subsequently establishing predetermined operating conditions for the synthesis of organic compounds in the presence of said finely divided contact material in a fluidized condition.

4. The starting-up procedure for the hydrogenation of a carbon oxide in the presence of a finely divided fluidized contact material comprising freshly reduced iron which comprises, prior to establishing predetermined operating conditions contacting said freshly reduced contact material with a gaseous mixture comprising hydrogen, passing said gaseous mixture through said finely divided contact material at a linear velocity sufilcient to maintain the contact material in a fluidized condition, maintaining an initial pressure of at least about 15 but below about 85 pounds per square inch gage, raising the pressure to at least about 100 pounds per square inch gage, contactlngat least 200 standard cubic feet of said gaseous mixture per pound of iron calculated as elementary metal and raising the temperature to at least about 600 F. during the period the pressure is raised from the initial pressure to about 100 pounds per square inch gage, maintaining the initial temperature of contact above about 450 F. and the maximum tem-' perature of contact below about 700 F. during the starting-up procedure, and subsequently establishing predetermined operating conditions for the synthesis of organic compounds in the presence of said finely divided contact material in a fluidized condition.

5. The process of claim 4 in which the contact material comprises a relatively low alkali iron catalyst and in which said initial temperature is at least about 500 F.

6. The process according to claim 4 in which said finely divided contact material comprises a relatively high alkali iron catalyst and in which said initial temperature is at least about 550 F.

7. The process according to claim 4 in which said predetermined operating conditions comprise a pressure below about 100 pounds per square inch gage.

8. The process according to claim 4 in which said fluidized condition of the contact material a is a fluid-bed condition.

9. The process according to claim 4 in which said predetermined operating conditions comprise a pressure above about pounds per square inch gage and the pressure is raised to said predetermined operating pressure at a rate equivalent to not greater than about 50 pounds per square inch per 100 standard cubic feet of said gaseous mixture.

10. A continuous process for the hydrogenation of carbon monoxide in the presence of a finely material in a fluidized condition, maintaining an 15 a gaseous efliuent comprising the desired organic products 01' the process and recovering same therefrom, removing a portion of the finely divided contact material from said first reaction zone and passing same to a second reaction zone, passing an oxidizing gas' through said finely divided contact material in said second reaction zone in a fluidized condition under conditions such that carbonaceous deposits are removed therefrom, removing a portion oi the contact material from said second reaction zone and passing same to a thirdreaction zone, passing a reducing gas through said contact material in a fluidized condition in said third reaction zone under conditionssuch that ferruginous material is reduced, removing contact material from said th rd reaction zone and passing same to a fourth reaction zone, in said fourth reaction zone pass-.

ing synthesis gas comprising hydrogen and carbon monoxide through said contact material in a fluidized condit on at an initial pressure of at least about but below about 90 pounds per square inch gage, raising the pressure in said fourth reaction zone to at least about 100 pounds.

per square inch gage, contacting at least about 200 standard cubic feet of synthesis gas per pound of iron calculated as elementary metal and establish ng a tem erature of at least about 600 F. during the period the pressure is raised from the initial pressure to about 100 pounds per square inch gage, maintaining the temperature of contactin said fourth reaction zone above 450 F. and below about 700 F., and returning the thus treated contact material from said fourth reaction zone to said first reaction zone.

11. A continuous process for the hydrogenation of carbon monoxide in the presence of a finely divided fluidized contact material comprising iron which comprises, in a first reaction zone continuously passing a synthesis gas compris ng hydrogen and carbon monoxide in a mol ratio between about 1:1 and about 4:1 through said contact material in a fluidized condition under condit ons such that organic compounds are produced, removing from said first reaction zone a gaseous effluent comprising the desired organic products of the process and recovering same therefrom, removing a portion of the contact material from said first reaction zone and passing same to a second react on zone, passing a reducing gas through said contact material in a fluidized condition in said second reaction zone under conditions such that ferruginous material s recluced, removing contact material from said second reaction zone and passing same to a third reaction zone, in said third reaction zone passing synthesis gas comprising hydrogen and carbon monoxide through said contact material in a fluidized condition at an initial pressure of at least about 15 but below about 90 pounds per square inch gage, raising the pressure in said third reaction zone to at least about 100 pounds per square inch gage, contacting at least about 200 standard cubic feet of synthesis gas per pound of iron calculated as elementary metal and establishing a temperature of at least about 600 F. during the per od the pressure is raised from the initial pressure to about 100 pounds per square inch gage, maintaining the temperature of contact in said third reaction zone above 450 F. and below about 700 F.. and rein a fluidized condition when used as such dur- I ing the hydrogenation of a carbon oxide which comprises reducing an oxide of iron to the elementary metal with hydrogen at an elevated temperature above about 600 F., contacting said reduced material in finely-divided form with a gaseous mixture of hydrogen and at least 15 volume per cent carbon monoxide at a linear gas velocity such that the finely-divided material is maintained in a fluidized condition and at an initial pressure of at least about 15 but below about pounds per square inch gage and at an initial temperature above about 450 F., raising the pressure to at least about 150 pounds per square inch gage and the temperature above about 600 F. during contact with said gaseous mixture, and contacting at least about 300 standard cubic feet of said gaseous mixture per pound 01' iron calculated as the elementary metal during the period the pressure is raised from the initial pressure to about 150 pounds per square inch gage. I

, 13. The catalyst claim 12.

14. An improved method for preparing an active iron catalyst capable of being maintained in a fluidized condition when used as such during the hydrogenation of a carbon oxide which prepared by the method of comprises reducing an oxide of iron with hydrogen at an elevated temperature, contacting said reduced material in a finely-divided form with a gaseous mixture of hydrogen and carbon monoxide at a linear gas velocity sufficient to maintain said finely-divided material in a fluidized condition and at an initial pressure of at least about 15 but below about 90 pounds per square inch gage and at an initial temperature above about 450 F., raising the pressure to at least about pounds per square inch gage and establishing a temperature of at least about 600 F. during contact with said gaseous mixture, and contacting at least about 200 standard cubic feet of said gaseous mixture per pound of iron calculated as the elementary metal during the period the pressure is raised from the initial pressure to about 100 pounds per square inch gage.

15. A method for preparing an active catalyst capable of being maintained in a fluidized condition when used as such during the hydrogenation of a carbon oxide which comprises fusing a solid material comprising naturally occurring" magnetite, substantially. completely reducing the fused magnetite to the elementary metal by contact with a gas comprising hydrogen at an elevated temperature, contacting the elementary metal thus produced in finely-divided form with a gaseous mixture comprising hydrogen and carbon monoxide at a linear gas velocity suiiicient to maintain said finely-divided material in a fluidized condition and at an initial temperature above about 450 F. and at an initial pressure of at least about 15 but below about 90 pounds per square inch gage, raising the pressure to at least about 100 pounds per square inch gage and establishing a temperature of at least about 600 F. during contact with said mixture of hydrogen and carbon monoxide, and contacting at least about 200 standard cubic feet of said mixture of hydrogen and carbon monoxide per pound of iron during the period the pressure is raised from the initial pressure to about 100 pounds per square inch gage. I

16. The catalyst prepared by the method of claim 15.

17. The starting-up procedure for the hydrogenation of a carbon oxide in the presence of 35 a finely-divided fluidized contact material comprising reduced iron which comprises, prior to establishing prede ermined operating conditions raised from the initial pressure to about 100' pounds per square inch gage raising the temperature above 600 F. and substantially above the predetermined operating temperature prior to establishing same and subsequently establishing predetermined operating conditions including temperature and pressure for the synthesis of organic compounds in the presence of said finelydivided contact material in a fluidized condition.

18. In a process for the hydrogenation of carbon monoxide in the presence of a finely-divided metallic iron catalyst containing an alkali as a promoter at predetermined operating conditions of a temperature between about 350 and about 700 F., a pressure between about 150 and about 600 pounds per square inch gage and a space velocity between about 100 and about 5000 volumes of total feed gas per hour per volume of catalyst to produce normally liquid organic compounds as products of the process, the method for pretreating the finely-divided metallic iron catalyst which comprises contacting the finely-divided metallic iron catalyst initially with a gaseous mixture comprising hydrogen and carbon monoxide at a temperature above about 500 F. and at a pressure of at least about 15 pounds per square inch gage but below about 90 pounds per square inch gage, passing said gaseous mixture through said finely-divided metallic iron catalyst at a velocity effective to maintain the catalyst in a fluidized condition, raising the pressure above about 150 pounds per square inch gage after contact with at least about 300 standard cubic feet of said gaseous mixture per pound of iron calculated as the elementary metal, raising the temperature above 600 F. and substantially initial pressure to about 150 pounds, and thereafter employing catalyst thus pretreated for the hydrogenation of carbon monoxide at the aforesaid predetermined operating conditions while 5 maintaining the catalyst in a fluidized condition.

19. In a process for the hydrogenation of carbon monoxide in the presence of a finely-divided metallic iron catalyst containing an alkali as a promoter, at predetermined operating condi m tions of a temperature between about 350 and about 700 F. and a pressure below about 100 pounds per square inch gage and at a space velocity between about 100 and about 5000 volumes of total feed gas per hour per volume of catalyst to produce normally liquid organic compounds as products of the process, the method for pretreating the finely-divided metallic iron catalyst which comprises contacting the finelydivided catalyst initially with a gaseous mixture comprising hydrogen and carbon monoxide at a temperature above about 500 F. and at a pressure of at least 15 pounds per square inch gage but below about 90 pounds per square inch gage, passing said gaseous mixture through said finelydivided catalyst at a, velocity effective to maintain the catalyst in a fluidized condition, raising the pressure above about 150 pounds per square inch gage after contact with at least about 300 standard cubic feet of said gaseous mixture per pound of iron calculated as the elementary metal, raising the temperature above about 600 F. and substantially above the predetermined operating temperature during the period the pressure is raised from the initial pressure to about 150 5 pounds, and thereafter employing catalyst thus pretreated for the hydrogenation of carbon monoxide at the aforesaid predetermined operating conditions while maintaining the catalyst in a fluidized condition.

-- HENRY G. MCGRATH.

REFERENCES CITED I The following references are of record in the file of this patent:

i UNITED STATES PATENTS Number Name Date 2,257,457 Fischer et al Sept. 30,1941 2,393,909 Johnson Jan. 29, 1946 3!: OTHER REFERENCES Fischer: Brennstafi-Chemie, vol. 16, page 2, Jan. 1, 1935. 

1. THE STARTING-UP PROCEDURE FOR THE HYDROGENATION OF CARBON MONOXIDE IN THE PRESENCE OF A FINELY DIVIDED FLUIDIZED CONTACT MATERIAL COMPRISING FUSED AND FRESHLY REDUCED IRON WHICH COMPRISES, PRIOR TO ESTABLISHING THE DESIRED OPERATING CONDITIONS INITIALLY CONTACTING SAID CONTACT MATERIAL COMPRISING FRESHLY REDUCED IRON WITH A GASEOUS MIXTURE HAVING A MOL RATION OF HYDROGEN TO CARBON MONOXIDE BETWEEN ABOUT 1:1 AND ABOUT 4:1 AT A PRESSURE OF AT LEAST ABOUT 1:1 AND ABOUT ABOUT 85 POUNDS PER SQUARE INCH GAGE AND AT A TEMPERATURE ABOVE ABOUT 450* F., PASSING SAID GASEOUS MIXTURE UPWARD THROUGH A MASS OF SAID FRESHLY REDUCED CONTACT MATERIAL AT A VELOCITY SUFFICIENT TO MAINTAIN THE CONTACT MATERIAL IN A FLUIDIZED CONDITION, GRADUALLY RAISING THE PRESSURE ABOVE ABOUT 100 POUNDS PER SQUARE INCH GAGE DURING CONTACT WITH AT LEAST ABOUT 200 STANDARD CUBIC FEET OF SYNTHESIS GAS PER POUND OF IRON CALCULATED AS THE ELEMENTARY METAL, DURING THE INITIAL CONTACT AT WHICH THE PRESSUR DOES NOT EXCEED ABOUT 100 POUNDS PER SQUARE INCH GAGE ESTABLISHING A TEMPERATURE OF AT LEAST ABOUT 600* F., AND SUBSEQUENTLY ESTABLISHING THE DESIRED OPERATING CONDITIONS FOR THE SYNTHESIS OF ORGANIC COMPOUND IN THE PRESENCE OF SAID FINELY-DIVIDED CONTACT MATERIAL IN A FLUIDIZED CONDITION. 